Desulfurization of asphaltene-containing hydrocarbonaceous black oils

ABSTRACT

Sulfurous, asphaltene-containing black oils are converted into desulfurized, lower-boiling hydrocarbon products. The process involves a combination of solvent deasphalting, thermal cracking and multiple-stage hydrocracking.

United States Patent 1191 Watkins Nov. 27, 1973 [54] DESULFURIZATION OF2.973313 2/1961 Pevere et al. 208 211 ASPHALTENE CONTAINING 3,287,254 IH1966 Paterson 208/86 3,637,483 1/1972 Carey 208/86 HYDROCARBONACEOUSBLACK OILS [75] Inventor: Charles H. Watkins, Arlington Helghts PrimaryExaminer-Herbert Levine [73] Assignee: Universal Oil Products Company,AttorneyJames R. l-loatson, Jr. et al.

Des Plaines, ll].

[22] Filed: Aug. 9, 1972 21 A l. N .2 279 124 l 1 PP 0 1 57 ABSTRACT[52] US. Cl 208/86, 208/61, 208/78,

208/80 Sulfurous, asphaltene-contammg black Oils are con- 51 Int. Cl clo13/00 vetted desulfurized 1Owar-boiling hydrocarbm [58] Field of Search208/86 61 78 8O Pmdws' The PMess inwlves a combinatim vent deasphalting,thermal cracking and multiple-stage [56] References Cited hydrocrackmg'UNITED STATES PATENTS 2,002,004 5/1935 Gard 208/14 10 Claims, 1 DrawingFigure Deapha/ring 7 t 7 I I l 1 1 5 Reactor Rear/o! 1 1 6 J 9 a t l ee/0 l 5 1 1 Solve/1! g Har Separafar 22 Iggy/rial C27 7 Vacuum Column /3z I F/acr/anarar 'DESULFURIZATION F ASPHALTENE-CONTAININGHYDROCARBONACEOUS BLACK OILS APPLICABILITY OF INVENTION The inventiondescribed herein is applicable to a process for the conversion ofpetroleum crude oil, and the heavier fractions derived therefrom, intodesulfurized, lower-boiling hydrocarbon products. More specifically, myinvention is directed toward a process for converting atmospheric towerbottoms, vacuum tower bottoms (vacuum residuum), crude oil residuum,topped crude oils, oils extracted from tar sands, etc., all of which arecommonly referred to in the art as black oils, and which contain asignificant quantity of asphaltic material and sulfurous compounds.

Petroleum crude oils, particularly the heavier oils extracted from tarsands, topped or reduced crudes, and vacuum residuum, contain highmolecular weight sulfurous compounds in exceedingly large quantities,nitrogenous compounds, high molecular weight organometallic complexes(principally containing nickel and vanadium) and lighthydrocarbon-insoluble material. The latter is principally found to becomplexed with sulfur and, to a significant extent, with the metalliccontaminants. In this regard, black oils differ considerably from heavygas oils which are not so severely contaminated, and which normally donot have as high a boiling range. In the petroleum refining art, a blackoil is generally characterized as a heavy hydrocarbonaceous material ofwhich more than about 10.0 percent (by volume) boils above a temperatureof 1,050F., and which has a gravity less than about 20.0 API. Sulfurconcentrations are exceedingly high, being more than 1.0 percent byweight, and often in excess of 3.0 percent by weight. There currentlyexists an abundant supply of such hydrocarbonaceous material; however,the utilization thereof, as a source of more valuable distillable liquidhydrocarbon products, is virtually precluded by present-day catalyticreaction techniques.

Knowledgeable experts are currently predicting a world-wide energycrisis in the not-too-distant future. Those possessing expertise in thefield of petroleum exploration, for example, are very much concernedwith the ever-dwindling reserve supply of natural gas in comparison tothe ever-increasing demand thereof. As a result of legislation beingimposed upon the sulfur content of fuel oils burned to meet certainenergy requirement, more and more energy suppliers are looking tonatural gas as a substitute. Several processes are presently beingproposed which, it is hoped, will alleviate the forthcoming criticalshortage of natural gas. These processes primarily involve theconversion of naphtha fractions, via steam reforming and shiftmethanation, into a synthetic natural gas. However, this in turn createsa shortage of naphtha boiling range material for ultimate utilization asmotor fuel, particularly with the advent of the need to produce "cleargasolines to avoid severe atmospheric pollution as the result ofmetal-containing motor fuel additives. Likewise, a shortage of keroseneboiling range fractions, principally employed as jet fuels, as well asgas oils, will stem from the necessity to convert such charge stocksinto suitable automotive motor fuel. A multitude of factors are,therefore, contributing to the developing energy crisis. Processingtechnology is required to insure the utilization of virtually 100.0percent of petroleum crude oil charge stocks. In the petroleum refiningart, this is commonly referred to as the bottom of the barrel.

The process encompassed by the present invention supplies at least someof the technology required to effect the catalytic conversion ofhydrocarbonaceous black oils into distillable hydrocarbons in volumetricyields exceeding 100.0 percent. Specific examples of the charge stocksto which the present scheme is adaptable, include a vacuum tower bottomshaving a gravity of 7.1 API, and containing 4.05 percent by weight ofsulfur and 23.7 percent by weight of asphaltenic material; a toppedcrude oil having a gravity of l 1.0 API, and containing 10.1 percent byweight of asphalts and 5.2 percent by weight of sulfur; a vacuumresiduum having a gravity of 8.8 API, and containing 3.0 percent byweight of sulfur and 4,300 ppm. of nitrogen; a vacuum bottoms having agravity of 5.4 API, and containing 6. 15 percent by wei ght of sulfur,233

weight of metallic contaminants and 12.8 percent by weight ofheptane-insoluble asphaltic material; and, a reduced crude having agravity of l 1.5 API, and containing 4.2 percent by weight of sulfur,3,400 ppm. of nitrogen, 166 ppm. of metals and 8.6 percent by weight ofasphaltenic material.

The paramount difficulty, heretofore encountered with fixed-bedcatalytic systems, has been the lack of catalyst stability whenprocessing at those conditions required to convert the sulfurouscompounds into hydrogen sulfide and hydrocarbons. At the operatingseverity required to achieve acceptable desulfurization, the asphalticmaterial, finally dispersed within the black oil, has the tendancy toflocculate and polymerize, and thus become deposited upon thecatalytically active surfaces of the catalyst. Furthermore, the metalliccontaminants filter into the internal cavities, or pores of the catalystand effectively shield active catalytic sites from the material beingprocessed. In addition to fixed-bed, vapor-phase hydrocracking, anotherattempted approach has been liquid-phase hydrogenation. In this type ofprocess, liquid-phase oil is passed upwardly, in admixture withhydrogen, into a fluidizedfixed bed of subdivided catalyst; althoughperhaps effective in removing at least a portion of theorganometalliccomplexes, this type process is ineffective with respect to theinsoluble asphalts. Since they are finally dispersed within the oil, theprobability of effecting simultaneous contact between the catalystparticle, the asphaltic material and the hydrogen required to preventcoke deposition is remote.

An integral part of the present combination process constitutes theremoval of asphaltic material prior to effecting the fixed-bed catalyticconversion of the charge stock. However, the separation of themetalcontaining asphaltic pitch is accomplished in a manner whichretains a convertible resin concentrate subsequently processed in amanner significantly increasing the volumetric yield of more valuabledistillable hydrocarbon products.

OBJECTS AND EMBODIMENTS A primary object of my invention is to provide aprocess for efiecting the conversion of hydrocarbonaceous black oils. Acorollary objective is to afford maximum yields of substantiallydesulfurized, distillable hydrocarbon products.

Another object of my invention is to increase the acceptable effectivelife of catalytic composites utilized in fixed-bed processing ofasphaltene-containing, sulfurous petroleum feed stocks.

Therefore, in one embodiment, my invention involves a process for theconversion of a sulfurous, asphaltene-containing hydrocarbonaceouscharge stock to produce lower-boiling, desulfurized hydrocarbonproducts, which process comprises the steps of: (a) deasphalting saidcharge stockwith a selective solvent, in a first solvent extractionzone, at extraction conditions selected to provide a solvent-leanasphaltic pitch and a solvent-rich, deasphalted first liquid phase; (b)deresining at least a portion of said first liquid phase with aselective solvent, in a second solvent extraction zone, at extractionconditions selected to provide a solvent-lean resin concentrate and asolvent-rich second liquid phase; reacting at least a portion of saidresin concentrate with hydrogen, in a catalytic first reaction zone, athydrocracking conditions selected to convert resins into lower-boilinghydrocarbons; (d) further reacting at least a portion of the resultingfirst reaction zone effluent in a non-catalytic second reaction zone, atthermal cracking conditions selected to produce additional lower-boilinghydrocarbons; (e) reacting at least a portion of the resulting thermallycracked product effluent and at least a portion of said second liquidphase, in a catalytic third reaction zone with hydrogen, athydrocracking conditions selected to produce additional lower-boilinghydrocarbons; and, (f) recovering said lower-boiling, desulfurizedhydrocarbon products from the resulting third reaction zone effluent.

Other objects and embodiments of my invention relate to additionaldetails regarding preferred catalytic ingredients, the concentration ofcomponents within the catalytic composites, preferred processingtechniques and similar particulars which are hereinafter given in thefollowing, more detailed summary of the present invention and thecombination process encompassed thereby. In one such other embodiment,the first reaction zone effluent is separated, in a first separationzone, at substantially the same temperature and pressure, to provide afirst principally vaporous phase and a third liquid phase, a portion ofsaid third liquid phase being reacted in said non-catalytic secondreaction zone. In another such embodiment, the second reaction zoneeffluent is separated, in a second separation zone, at substantially thesame temperature and at a reduced pressure in the range of fromsubatmospheric to about 200 psig., to provide a second principallyvaporous phase and a heavy resin by-product.

In still another embodiment, the extraction conditions in said secondextraction zone include a higher temperature than that in said firstextraction zone.

SUMMARY OF INVENTION The combination process, encompassed by the presentinventive concept, utilizes two solvent extraction zones for the purposeof (l) deasphalting the charge stock to remove an asphaltic pitch and(2) deresining the deasphalted oil to recover a resin concentrate, andto provide a deresined, deasphalted oil. A first fixedbed catalyticreaction zone is utilized to process the resin concentrate to producelower-boiling hydrocarbon products therefrom. At least a portion of theresin concentrate product effluent is subjected to thermal cracking,after which heavy resins are removed as a byproduct and the remainder iscombined with the deasphalted, deresined oil resulting from the secondsolvent extraction zone. This mixture is then subjected to hydrocrackingin a second fixed-bed catalytic reaction zone in order to produceadditional lower-boiling hydrocarbon products. Desulfurization iseffected to some extent in the first fixed-bed catalytic reaction zone,and is completed in the second fixed-bed reaction zone to the extentthat the normally liquid hydrocarbon products are substantiallysulfur-free. The inventive concept, upon which the present combinationprocess is founded, stems from recognition that black oils, of the typehereinbefore described, contain a convertible resin concentrate inaddition to an asphaltic fraction. Therefore, the first solventextraction zone functions to reject the asphaltic pitch whilemaintaining the resin concentrate in the solvent-rich deasphalted oil(DAO) phase. The latter is subjected to solvent extraction, preferablyat a higher temperature, to recover separately the resin concentrate,and to provide a deresined oil (DRO). In this manner, the asphalticpitch is removed from the charge stock prior to processing over thefixed-bed catalytic composites, and the resin concentrate can beprocessed at an operating severity required to produce lower-boilinghydrocarbons without incurring adverse effects with respect to theremainder of the deasphalted oil. By the same token, the deasphalted oilcan be processed at conditions conducive to the production of normallyliquid hydrocarbons in the absence of adverse effects stemming from thepresence of the resinous material. As hereinafter set forth, the resinconcentrate will generally be processed at a lower severity level thanthat imposed on the reaction system processing the deasphalted oil inadmixture with other internally produced streams, the sources of whichare hereinafter set forth.

The present combination process utilizes two solvent extraction zones toprecipitate an asphaltic concentrate and to provide a resin concentratefor subsequent processing. It must necessarily be acknowledged that theprior art is replete with a wide spectrum of techniques employed foreffecting solvent deasphalting of asphaltene-containinghydrocarbonaceous charge stocks. It is understood, therefore, that noattempt is herein made to claim solvent deasphalting other than as it isemployed as an integral element of the present combination process. Anysuitable solvent deasphalting technique known in the prior art may beemployed, several examples of which are to be found in the referenceshereinafter briefly described. In the interest of brevity, no attemptwill be made to delineate exhaustively the solvent deasphalting art.

Exemplary of such prior art is U. S. Pat. No. 1,948,296 (Class 2084)which discloses a process for obtaining a particularly good road-typeasphalt product. The separated asphaltic fraction is admixed with asuitable oil (lubricating oil, gas oil, etc.) and subjected tooxidation. For effecting the precipitation of the asphaltic fraction,suitable solvents are described as including light petroleum hydrocarbonmixtures such as naphtha, light petroleum fractions comprising propane,n-butane and isobutane, certain alcohols, ether and mixtures thereof,etc. In U. S. Pat. No. 2,101,308 (Class 208-309) similar solvents areutilized to precipitate only a portion of the asphaltic fraction. Asolvent of an altogether different character, for example liquid sulfurdioxide, is utilized to separate the resulting exhydrocarbons such aspropane, n-butane, isobutane, as

well as ethane, ethylene, propylene, n-butylene, isobu- 'tylene,pentane, isopentane and mixtures thereof.

Identical solvents, utilized to precipitate asphaltics, are disclosed inUS. Pat. No. 2,587,643 (Class 2084,09). These are, however utilized inadmixture with an organic carbonate.

Conspicuously absent from such prior art is a recognition of thedifference between the asphaltic fraction and the convertible resinconcentrate. Certainly there is found no awareness that the resinconcentrate can be processed in a fixed-bed catalytic reaction system toproduce more valuable distillable hydrocarbons of lower sulfur content.The multiple solvent extraction zones of the present combination processrecover a convertible resin concetrate as an essential feature of thepresent invention, and the prior are is silent with respect to thisoperating technique.

Although both solvent extraction zones will function at generally thesame operating conditions, a preferred technique involves precipitatingthe asphaltic pitch in the first solvent extraction zone at a lowertemperature than is utilized to rec over a resin concentrate in thesecond solvent extraction zone. Suitable extraction conditions include atemperature in the range of about 50F. to about 600F., and preferablyfrom about 100F. to about 400F.; the pressure will be maintained withinthe range of about 100 to about 1,000 psig., and

' preferably from about 200 to about 600 psig. The precise operatingconditions will generally depend upon the physical characteristics ofthe charge stock as well as the selected solvent. In general thetemperature and pressure are selected to maintain the solvent extractionoperations in liquid phase and, with respect to thefirst extractionzone, to insurethat substantially all the asphaltic pitch is removed inthe solvent-lean heavy phase with the resin concentrate being retainedin the solvent-rich deasphalted oil phase. Suitable solvents, forutilization in the present combination process, include thosehereinbefore described with respect to prior art deasphaltingtechniques. Thus, it is contemplated that the solvent will be selectedfrom the group of light hydrocarbons such asethane, methane, propane,butane, isobutane, pentane, isopentane, neopentane, hexane, isohexane,heptane, the monoolefinic counter-parts thereof, etc. Furthermore, thesolvent may be a normally liquid naphtha fraction containinghydrocarbons having from about five to about 14 carbon atoms permolecule, and preferably a naphtha fraction having an end boiling pointbelow about 200 F. The solvent-rich normally liquid phase is generallyintroduced into a suitable solvent recovery system,

the design and techniques of which are thoroughly described in the priorart.

The terms, deasphalting and deresining, as employed in thisspecification and the appended claims, connote the rejection of anasphaltic pitch and, subsequently, a resin concentrate. The precisechemical and physical nature of these two fractions is largely dependentupon the origin of the crude oil and the conditions utilized in theextraction zones, the latter including the nature of the selectivesolvent. As currently practiced, solvent deasphalting generally refersto a one-stage precipitation operation as applied to anasphalt-containing residuum, whereas deresining, as in the presentprocess, refers to a similar treatment performed on an essentiallyasphalt-free residuum. Regardless of their precise nature, deasphaltingand deresining apply to the rejection of two contiguous bottomsfractions.

In accordance with the present invention, the deresining operationinvolves the use of a greater solvent/oil volumetric ratio and a highertemperature. For example, in propane deasphalting, a solvent/oil ratioof about 60:10 and a temperatuere of F. to F. are typical. For thederesining operation, a propane/oil ratio of about l0.0:l.0 and atemperature of about 140F. to about F. are typical.

The asphaltic pitch, obtained from the deasphalting operation willexhibit an average molecular weight in the range of 3,000 to 6,000, orhigher, and willcontain from 75.0 percent to 90.0 percent of the metalspresent in the fresh feed charge stock. The sulfur content will beapproximately twice that of the fresh charge. The molecular weight ofthe resin concentrate will be lower, and in the range of about 1,000 toabout 4,000; it will contain only a minor portion of the virgin metalcontaminants, and have a sulfur content about one and one-half timesgreater than the feed stock.

before describing my invention further, and especially with respect tothe embodiment illustrated in the accompanying drawing, severaldefinitions are believed necessary in order that a clear understandingbe obtained. In the present specification and appended claims, apressure substantially the same as, of a temperature substantially thesame as," is intended to connote the pressure or temperature on adownstream vessel, allowing only for the normal pressure drop due tofluid flow through the system, and the normal temperature loss due tothe transfer of material from one zone to another. For example, wherethe pressure at the inlet to the first catalytic reaction zone is about3,000 psig., and the effluent temperature is about 750F., the firstseparation zone will function at substantially the same pressure andtemperature of about 2,925 psig. and 750F. Similarly, the utilization ofthe phrase at a least a portion of," when referringtoeither aprincipally vaporous phase, or a principallydiquid phase, is intended toencompass both an aliquot portion as well as a select fraction. Thus, atleast a portion of a hydrogen-rich principally vaporous phase isrecycled to a catalytic reaction zone, following the removal of hydrogensulfide therefrom, while at least a portion of the liquid phase (in thiscase an aliquot portion) may be recycled to a catalytic reaction zonetocombine with the fresh feed charge stock thereto.

As previously set forth, the combination process of the presentinvention utilizes two hydrocracking reaction zones. In most instances,the catalytic composites disposed within the two reaction zones will beof different physical and chemical characteristics; it is understood,however, that they may be identical. Regardless, the catalyticcomposites comprise metallic components selectedfrom the metals ofGroups Vl-B and Vlll of the Periodic Table, and compounds thereof. Thus,in accordance with the Periodic Table of The Elements, E. H. Sargent andCo., 1964, suitable metallic components are those selected from thegroup consisting of chromium, molybdenum, tungsten, iron, ruthenium,osmium, cobalt, rhodium, iridium, nickel, palladium and platinum.Additionally, recent investigations have indicated that catalyticcomposites, for utilization with excessively high-sulfur content feedstocks, are improved through the incorporation of a zinc, tin and/orbismuth component. Throughout the present specification and the appendedclaims, the use of the term component, when referring to thecatalytically active metal, is intended to connote the existence of themetal within the catalytic composite either in some combined form, or inthe elemental state. Regardless, the stated concentration of themetallic component is computed on the basis of the elemental metal.While neither the precise composition, nor the method of manufacturingthe catalytic composites, is considered essential to my invention,certain aspects are preferred. For example, since the charge stock tothe present process is of a high-boiling nature, it is preferred thatthe metallic components of the catalyst possess the propensity foreffecting hydrocracking while simultaneously promoting the conversion ofsulfurous compounds into hydrogen sulfide and hydrocarbons. Theconcentration of the catalytically active metallic component, orcomponents, is primarily dependent upon the particular metal as well asthe physical and/or chemical characteristics of the feed stock. Forexample, the metallic components of Group VI-B are generally present inan amount within the range of about 4.0 percent to about 30.0 percent byweight, the iron-group metals in an amount within the range of about 0.2percent to about 10.0 percent by weight, whereas the moble metals ofGroup VIII are preferably present in an amount within the range of about0.1 percent by weight, all of which are calculated as if thesecomponents existed within the catalytic composite in the elementalstate. When a zinc, tin and/or bismuth component is utilized, the samewill be present in an amount of about 0.1 percent to about 5.0 percentby weight.

The porous carrier material, with which the catalytically activemetallic components are combined, is a refractory inorganic oxide of thecharacter thoroughly described in the literature. When of the amorphoustype, alumina, or alumina in combination with about 10.0 percent toabout 90.0 percent by weight of silica is preferred. It is oftenappropriate to utilize a carrier material comprising a crystallinealuminosilicate, or zeolitic molecular sieve. ln most instances, such acarrier material will be utilized in processing the deasphalted oil inthe second catalytic reaction zone. The zeolitic material includesmordenite, faujasite, Type A or Type U molecular sieves, etc. These maybe employed in a substantially pure state; however, it is contemplatedthat the zeolitic material may be included within an amorphous matrixsuch as silica, alumina, and mixtures of alumina and silica.

It is further contemplated that a halogen component may be combined withthe other components of the catalytic composite. Although the preciseform of the chemistry of association of the halogen components with thecarrier material and metallic components is not accurately known, it iscustomary in the art to refer to the halogen component as being combinedwith the carrier material or with the other ingredients of the catalyst.The halogen may be either fluorine, chlorine, iodine, bromine, ormixtures thereof, with fluorine and chlorine being particularlypreferred. The quantity of halogen is such that the final catalyticcomposite contains about 0.1 percent to about 3.5 percent by weight, andpreferably from about 0.5 percent to about 1.5 percent by weight,calculated on the basis of the elemental halogen.

The metallic components may be incorporated within the catalyticcomposite in any suitable manner including co-precipitation orcogellation with the carrier material, ion-exchange or impregnation ofthe carrier material. Following the incorporation of the metalliccomponents, the catalyst is dried and subjected to a high temperaturecalcination or oxidation technique at a temperature of about 750F. toabout 1,300F. When a crystalline aluminosilicate is utilized as part ofthe carrier material, the upper limit for the calcination technique ispreferably about 1,000F.

With respect to the catalyst utilized in the catalytic first reactionzone, a preferred composite is of the character described in US. Pat.No. 3,640,817 (Class 208-59). Briefly, this catalyst consists of acarrier material of alumina and silica containing from about 5.0 percentto about 30.0 percent by weight of boron phosphate, and has more thanabout 50.0 percent of its macropore volume consisting of pores havingnominal diameters greater than about 1,000 Angstroms.

Prior to its utilization for the desulfurization/hydrocracking ofhydrocarbons, the dried and calcined catalytic composite may besubjected to a substantially water-free reduction technique.Substantially pure and dry hydrogen (less than about 30.0 volumetn'cppm. of water) is employed as the reducing agent. The calcined compositeis contacted at a temperature of about 800F. to about 1,200F. and for aperiod of about 0.5 to about 10 hours. This reduction technique may beperformed in situ prior to introducing the charge stock.

Additional improvements are generally obtained when the reducedcomposite is subjected to presulfiding for the purpose of incorporatingtherewith from about 0.05 percent to about 0.5 percent by weight ofsulfur, on an elemental basis. The presulfiding treatment is effected inthe presence of hydrogen and a suitable sulfur-containing compound suchas hydrogen sulfide, a low molecular weight mercaptan, various organicsulfides, carbon disulfide, etc. One technique involves treating thereduced catalyst with a sulfiding gas, such as a mixture of hydrogen andhydrogen sulfide having about 10 moles of hydrogen per mole of hydrogensulfide, and at conditions selected to effect the desired incorporationof sulfur. Presulfiding may also be effected in situ by way of charginga relatively low boiling hydrocarbon feed containing sulfurouscompounds.

As hereinbefore set forth, the present invention utilizes two fixed-bedcatalytic reaction zones and a noncatalytic thermal cracking zone. Theresin concentrate is processed in a catalytic first reaction zone, andthe product effluent therefrom is separated in a hot separator atsubstantially the same pressure. The principal function served by thehot separator is to separate the mixed-phase product effluent into avapor phase rich in hydrogen and a principally liquid phase which maycontain from about 10.0 mol.% to about 40.0 mol.% of dissolved hydrogen.In a preferred embodiment, the total reaction product effluent from thecatalytic first reaction zone is utilized as a heat-exchange medium inorder to lower thetemperature thereof to a level in the range of about700F. to about 800F. The liquid phase from the hot separator may berecycled, at least in part, to combine with thefresh resin concentrate,thereby serving as a diluent for the heavier constituents thereof. Thequantity of the liquid phase diverted in this manner is such that thecombined feed ratio to the catalytic first reaction zone, being definedas total volumes of liquid charge per volume of fresh liquid charge, iswithin the range of about 1.1:1 to about 3.5: 1 The remaining portion ofthe principally liquid phase from the hot separator is introduced intothe thermal cracking reaction zone, or coil, at a reduced pressure inthe range of about 200 psig. to about 500 psig. and at a temperature offrom about 700F. to about950 F. As hereinafter indicated in thedescription of the accompanying drawing, the product effluent from thethermal cracking coil is introduced into a vacuum flash columnmaintained at about to about 60 mm. Hg., absolute. The principalfunction of the vacuum flash zone is to concentrate the remaining heavy,metal-containing resins as a byproduct stream while recoveringdistillable hydrocarbons as a principally vaporous phase. The vaporousphase from the vacuum flash zone, in combination with the vaporous phaserecovered from the hot separator and the deasphalted andderesined oil isprocessed in the catalytic third reaction zone toproduce additionaldesulfurized lower-boiling hydrocarbon products.

With respect tothe two catalytic reaction zones, the operatingconditions of temperature, pressure, liquid hourly space velocity andhydrogen/hydrocarbon ratio will be within the same ranges. However, apreferred technique dictates operating the catalytic first reactionzone, processing the resin concentrate, at a lower severity than thatimposed upon the catalytic third reac tionzone. The variance inoperating severity levels between the two catalytic reaction zones isreadily obtained through the adjustment of the pressure, maximumcatalyst bed temperature and liquid hourly space velocity. the higherseverity operation will normally be effected at an increased pressure,an increased maximum catalyst bed temperature and at a decreased liquidhourly space velocity, or. some combination thereof. The maximumcatalyst bed temperature within the catalytic first reaction zone willbe at least about 20F. lower than that maintained within the catalyticthird reaction zone, in most instances.

With respect to the operating conditions impose upon the catalyticreaction zones, they are selected primarily to effect the conversion ofsulfurous compounds into hydrogen sulfide and hydrocarbons, whilesimultaneously inducing hydrocracking reactions to produce lower-boilinghydrocarbon products. As hereinbefore set forth, the operatingconditions imposed upon the catalytic third reaction zone will result ina higher operating severity. Suitable ranges for the various variableswill generally be the same for both reaction systems. Thus, the pressurewill range from about 500 to about 3,500 psig., and preferably fromabout 500 to about 2,500 psig. The maximum catalyst bed temperature willbe within the range of about 600F. to about 900F. In view of the factthat the reactions being effected in the catalytic reaction zones areprincipally exothermic, an increasing temperature gradient will beexperienced as the reactants traverse the catalyst bed. judiciousoperating techniques dictate that the increasing temperature gradient belimited to a maximum of about 100F., and, in order to control theincreasing temperature gradient, it is within the scope of the presentinvention to employ quench streams, either normally liquid, or normallygaseous, introduced at one or more intermediate loci of the catalystbed. The hydrogen concentration is expressed as scf./Bbl. of charge, andwill usually be within the range of about 1,000 to about 30,000. Liquidhourly space velocities, defined as volumes of normally liquidhydrocarbons charged per hour, per volume of catalyst disposed withinthe reaction zone, will be from about 0.25 to about 2.50. In addition tothe temperature variable, the liquid hourly spaced velocity isconveniently utilized to adjust the operating severity between the twocatalytic reaction zones. Thus, the liquid hourly space velocity throughthe second reaction zone will generally be less than that through thefirst reaction zone.

That portion of the effluent from the catalytic third reaction zoneboiling at a temperature above that de sired with respect to therecovered product streams, may be recycled in order to produceadditional lowerboiling hydrocarbon products. When this technique isutilized, the combined feed ratio,-defined as total volumes of normallyliquid charge to the catalytic third reaction zone, per volume of freshcharge thereto, will be within the range of about 1.121 to about 6.0:1.

Other conditions and preferred operating techniques will be given inconjunction with the following description of the present process. Infurther describing this process, reference will be made to theaccompanying figure which illustrates one specific embodiment. In thedrawing, the embodiment is presented by means of a simplified flowdiagram in which many details such as pumps, instrumentation andcontrols, heat-exchange and heat-recovery circuits, valving, start-uplines and similar hardware have been omitted as being nonessential to anunderstanding of the techniques involved. The use of suchmiscellaneousappurtenances,

to modify the process, are well within the purview of one skilled in theart. I

The major vessels integrated within the combination process of thepresent invention, as illustrated in the drawing, are as follows: thefirst solvent extraction zone is deasphalting zone 2, while the secondsolvent extraction zone is deresining zone 5; the catalytic firstreaction zone is reactor 9, the non-catalytic, second reaction reactionzone is thermal coil 14 and the catalytic third reaction zone is reactor20; and, the first separation zone is hot separator 11, the secondseparation zone is vacuum column 16 and the third separation zone iscold separator 22. A fourth separation zone is illustrated asfractionator 24 and functions to recover the various desired productfractions. For example, a propane-minus stream may be recovered throughline 25, a butane concentrate through line 26, a combined pentane/hexaneconcentrate in line 27, a naphtha boiling range, heptane-400F. productin line 28 and 400F.650F. middle-distillate through line 29.

DESCRIPTION OF DRAWING The accompanying drawing will be described inconjunction with a commercially scaled unit designed to process 80,000Bbl./day of vacuum column bottoms. Charge stock analyses indicate agravity of about 10.1 API, 3.08 percent by weight of sulfur, 186 weightppm. of metals, a Conradson Carbon content of 15.8 percent and aheptane-insoluble portion in the amount of 5.2 percent by weight. Thedesired product slate includes a light naphtha (heptane-275F.), a heavynaphtha (275F380F.) and a diesel fuel (380F-650F.). All the desiredfractions are intended to be substantially free from nitrogenous andsulfurous compounds. In the description, the yields, unless otherwisespecifically stated, are given in weight percent, and are based upon thevacuum bottoms charge.

The charge stock is introduced, via line 1, into a deasphalting zone 2,wherein it countercurrently contacts a pentane/butane solvent introducedvia line 31. The solvent extraction is effected in substantially liquidphase at a pressure of about 400 psig. and a temperature of 245F., witha solvent/oil volumetric ratio of 30:10 A solvent-lean asphaltic pitch,in the amount of about 15.0 percent by weight, having a gravity of 8.1APl, is withdrawn through line 3 while a solventrich, resin-containingfirst liquid phase is recovered via line 4.

The first liquid phase countercurrently contacts addi tionalpentane-butane solvent, at a solvent/oil volumetric ratio of 5.0210,introduced into deresining zone 5 by way of line 32. The temperature is300F. and the pressure about 400 psig., which produces a resinconcentrate in the amount of about 29.0 percent by weight, having agravity of 1.62 API, precipitated and withdrawn by way of line 6. Theresin concentrate has a metals concentration of 136 ppm. by weight. Asolvent-rich second liquid phase is removed through line 7, in an amountof 56.0 percent by weight, and is subsequently reacted with hydrogen inreactor 20. The deresined oil in line 7 has a gravity of 16.9 API,and'contains 2.0 percent by weight of sulfur and only 3.0 ppm. of metalcontaminants. The 29.0 percent by weight of resin concentrate continuesthrough line 6, is admixed with a hydrogen-rich, principally vaporousphase from line 8, and introduced thereby into a catalytic firstreaction zone 9.

The hydrogen concentration in reactor 9 is about 5,000 scf./Bbl. and thepressure is maintained at 2,500 psig. A temperature gradient of 100F. iscontrolled through the use of a hydrogen quench stream, while thereactants traverse the catalyst bed at a liquid hourly space velocity of1.0, to result in a maximum catalyst bed temperature of 875F. Reactor 9contains a catalyst of 1.89 percent by weight of nickel, 16.0 percentmolybdenum, 8.78 percent of boron phosphate, 6.97 percent silica and66.96 percent by weight of alumina. The first reaction zone producteffluent is withdrawn through line 10, and introduced into hot separator11 at substantially the same pressure and a temperature of about 750F. Afirst principally vaporous phase is withdrawn through line 12, and isadmixed with the deasphalted and deresined oil in line 7. A heavy, thirdprincipally liquid phase is removed via line 13, to be charged tothermal coil 14. A material balance around.

hot separator 11 is given in the following Table I:

TABLE I: Hot Separator Balance Line No. l0 l2 l3 Gases 1.5 1.5 TraceLight Naphtha 0.3 0.3 Trace Heavy Naphtha 0.8 0.8 Trace Diesel Fuel 554.0 1.5 Heavy Oil 6.9 0.9 6.0 Resins 14.5 Trace 14.5 TOTALS: 29.5 7.522.0

The figures presented in Table l are inclusive of hydrogen consumptionin an amount of 0.5 percent by weight.

The remaining 14.5 percent by weight of resins are introduced into anon-catalytic, second reaction zone 14 (thermal coil) at a pressure ofabout 200 psig. Thermal reactions therein are carried out at atemperature of 950F., to produce a thermally cracked product effluent inline 15 which is separated in vacuum column 16 at a pressure of 60 mm.Hg, absolute and a temperature of 800F. Distillable hydrocarbons, as thesecond principally vaporous phase, are removed through line 17, to becombined with the derisined oil in line 7, and a metal-containing, heavyresin by-product is withdrawn from the process by way of line 18, in anamount of 4.5 percent by weight. It should be noted that more than 86.0percent of the resin concentrate, precipitated in deresining zone 5, hasbeen converted into more valuable, distillable hydrocarbon products. Thematerial balance around vacuum column 16 is presented in the followingTable 11:

TABLE I1: Vacuum Column Balance Line No. l5 l7 l8 Gases 0.5 0.5 LightNaphtha 0.4 0.4 Heavy Naphtha 1.0 1.0 Diesel Fuel 5.5 5.5 Heavy Oil 10.110.1 Heavy Resins 4.5 4.5

The deasphalted oil in line 7, in admixture with the first vaporousphase from hot separator 11 (line 12), the second vaporous phase fromvacuum column 16 (line 17) and a heavy oil recycle stream from line 30,the source of which is hereafter set forth, is introduced into catalyticthird reaction zone 20 in admixture with a hydrogen-rich vaporous phasein line 19.

Reactor 20 has disposed therein a catalyst of 1.9 percent by weight ofnickel, 14.1 percent by weight of molybdenum, 27.3 percent of silica and56.7 percent alumina. The hydrogen concentration is about 6,000scf./Bbl., and the pressure is maintained at about 2,400 psig. Thenormally liquid portion of the feed stock traverses the catalyst bed ata liquid hourly space velocity of 0.6, and the maximum catalyst bedtemperature is controlled at about 875F.

The resulting reaction product effluent passes through line 21, atsubstantially the same pressure, and, after being used as aheat-exchange medium and further cooling, into cold seaparator 22 at atemperature of about F. A hydrogen-rich, third vaporous phase iswithdrawn via line 8, and in part recycled thereby to reactor 9; aportion is diverted from line 8 by line 19, as hydrogen recycle toreactor 20. The normally liquid portion of the product effluent isremoved by way of line 23 and introduced thereby into fractionator 24for separation into the various product streams. A butaneminus stream isrecovered via line 25, a pentane/hexane concentrate through line 26, thelight naphtha fraction via line 27, the heavy naphtha fraction throughline 28 and the diesel oil by way of line 29. Heavy oil, boiling beyondthe diesel oil end boiling point of 650F., is recycled in an amount of20.0 percent by weight, through lines 30, 17 and 7, for furtherconversion in reactor 20. The balance around reactor 20, including theheavy oil recycle, is presented in the following Table III:

TABLE I11: Reactor 20 Balance Line No. 7 (Charge) 23 (Product)* Gases2.0 7.5 Light Naphtha 0.7 3.0 Heavy Naphtha 1.8 10.5 Diesel Fuel 9.561.5 Heavy Oil 31.0 20.0 PDR 56.0

* Includes total hydrogen consumption of 2.0% by weight Propanederesined oil from extraction zone Overall product yield and componentdistributions are presented in the following Table IV, and are based ona fresh feed charge stock rate of 80,000 Bbl./day.:

TABLE IV: Product Yield and Distribution Component Wt.% Vol.% BbL/dayAmmonia 0.2 Hydrogen Sulfide 2.2 Methane 0.2 Ethane 0.4 Propane 0.7Butanes 1.3 2.3 1,840 Pentanes 1.2 1.9 1,5 Hexanes 1.3 1.9 1,520Heptane-275F. 3.0 4.2 3,360 275F.380F. 10.5 13.8 11,040 380F.-650F. 61.574.1 59,280 Asphaltic Pitch 15.0 13.3 10,640 Heavy Resins 4.5 4.1 3,280

TOTALS: 102.0 1 15.6 92,480

All the normally liquid streams, including the pentane/hexaneconcentrate, which may be supplied to an isomerization zone to producehigh octane isomers, indicate substantially no sulfurand/ornitrogencontaining compounds. The propane and butanes may be recoveredas a concentrate and employed as the feed stream to a steam reformingunit to produce a methane-rich synthetic natural gas, or as the feed toa dehydrogenation unit to produce olefins for subsequent alkylation to ahigh octane alkylate motor fuel.

The foregoing indicates the method of effecting the present combinationprocess and the benefits afforded through the utilization thereof.

I claim as my invention:

1. A process for the conversion of a sulfurous, asphaltene-containinghydrocarbonaceous charge stock to produce lower-boiling, desulfurizedhydrocarbon products, which process comprises the steps of:

a. deasphalting said charge stock with a selective solvent, in a firstsolvent extraction zone, at extraction conditions selected to provide asolvent-lean asphaltic pitch and a solvent-rich, deasphalted firstliquid phase;

b. deresining at least a portion of said first liquid phase with aselective solvent, in a second solvent extraction zone, at extractionconditions selected to provide a solvent-lean resin concentrate and asolvent-rich second liquid phase;

c. reacting at least a portion of said resin concentrate with hydrogen,in a catalytic first reaction zone, at hydrocracking conditions selectedto convert resins into lower-boiling hydrocarbons;

(1. further reacting at least a portion of the resulting first reactionzone effluent in a non-catalytic second reaction zone, at thennalcracking conditions selected to produce additional lower-boilinghydrocarbons;

e. reacting at least a portion of the resulting thermally crackedproduct effluent and at least a portion of said second liquid phase, ina catalytic third reaction zone, with hydrogen, at hydrocrackingconditions selected to produce additional lower-boiling hydrocarbons;and,

f. recovering said lower-boiling, desulfurized hydrocarbon products fromthe resulting third reaction zone effluent.

2. The process of claim 1 further characterized in that said firstreaction zone efiluent is separated, in a first separation zone, atsubstantially the same temperature and pressure, to provide a firstprincipally vaporous phase and a third liquid phase, and reacting atleast a portion of said third liquid phase in said non-catalytic secondreaction zone.

3. The process of claim 2 further characterized in that at least aportion of said first vaporous phase is reacted with hydrogen in saidthird reaction zone.

4. The process of claim 1 further characterized in that said secondreaction zone effluent is separated, in a second separation zone, atsubstantially the same temperature and at a reduced pressure in therange of from subatmospheric to about 200 psig, to provide a secondprincipally vaporous phase and a heavy resin product.

5. The process of claim 4 further characterized in that at least aportion of said second vaporous phase is reacted with hydrogen in saidthird reaction zone.

6. The process of claim 1 further characterized in that said thirdreaction zone effluent is separated, in a third separation zone, atsubstantially the same pressure and at a temperature in the range ofabout 60F. to about F., to provide a fourth liquid phase and ahydrogen-rich third principally vaporous phase.

7. The process of claim 6 further characterized in that at least aportion of said third vaporous phase is recycled to said first reactionzone.

8. The process of claim 6 further characterized in that at least aportion of said fourth liquid phase is recycled to said third reactionzone.

9. The process of claim 1 further characterized in that the extractionconditions in said second extraction zone include a higher temperaturethan that in said first extraction zone.

10. The process of claim 1 further characterized in that said portion ofthe resulting first reaction zone effluent is reacted with hydrogen insaid non-catalytic second reaction zone.

2. The process of claim 1 further characterized in that said firstreaction zone effluent is separated, in a first separation zone, atsubstantially the same temperature and pressure, to provide a firstprincipally vaporous phase and a third liquid phase, and reacting atleast a portion of said third liquid phase in said non-catalytic secondreaction zone.
 3. The process of claim 2 further characterized in thatat least a portion of said first vaporous phase is reacted with hydrogenin said third reaction zone.
 4. The process of claim 1 furthercharacterized in that said second reaction zone effluent is separated,in a second separation zone, at substantially the same temperature andat a reduced pressure in the range of from subatmospheric to about 200psig., to provide a second principally vaporous phase and a heavy resinproduct.
 5. The process of claim 4 further characterized in that atleast a portion of said second vaporous phase is reacted with hydrogenin said third reaction zone.
 6. The process of claim 1 furthercharacterized in that said third reaction zone effluent is separated, ina third separation zone, at substantially the same pressure and at atemperature in the range of about 60*F. to about 140*F., to provide afourth liquid phase and a hydrogen-rich third principally vaporousphase.
 7. The process of claim 6 further characterized in that at leasta portion of said third vaporous phase is recycled to said firstreaction zone.
 8. The process of claim 6 further characterized in thatat least a portion of said fourth liquid phase is recycled to said thirdreaction zone.
 9. The process of claim 1 further characterized in thatthe extraction conditions in said second extraction zone include ahigher temperature than that in said first extraction zone.
 10. Theprocess of claim 1 further characterized in that said portion of theresulting first reaction zone effluent is reacted with hydrogen in saidnon-catalytic second reaction zone.